Serial reforming with platinum on a non acidic support and platinum-rhenium on an acidic support

ABSTRACT

A PROCESS FOR THE CATALYTIC REFORMING OF NAPHTHENE AND PARAFFIN-CONTAINING HYDROCARBONS TO IMPROVE THEIR OCTANE RATING IS DISCLOSED WHICH INVOLVES THE USE OF A SERIES OF CATALYTIC REACTION ZONES OF WHICH ONE OF THE FIRST CONTAINS A SUPPORTED PLATINUM GROUP METAL-CONTAINING, LOW ACIDITY CATALYST WHICH IS SUBSTANTIALLY DEVOID OF RHENIUM AND WHICH SERVES TO DEHYDROGENATE NAPHTHENES, AND ONE OF THE LAST OF WHICH ZONES CONTAINS A SUPPORTED PLATINUM GROUP METAL AND RHENIUM-CONTAINING CATALYST OF HIGHER ACIDITY WHICH SERVES TO DEHYDROCYCLIZE PARAFFINS.

Int. (:1. clo 35/08, 39/00 US. Cl. 208-65 24 Claims ABSTRACT OF THEDISCLOSURE A process for the catalytic reforming of naphthene andparaflin-containing hydrocarbons to improve their octane rating isdisclosed which involves the use of a series of catalytic reaction zonesof which one of the first contains a supported platinum groupmetal-containing, low acidity catalyst which is substantially devoid ofrhenium and which serves to dehydrogenate naphthenes, and one of thelast of which zones contains a supported platinum group metal andrhenium-containing catalyst of higher acidity which serves todehydrocyclize paraifins.

The present invention is directed to the catalytic reforming of gasolineboiling range hydrocarbons. More particularly, this invention isconcerned with the catalytic reforming in the presence of molecularhydrogen of naphthene and paraflin-containing hydrocarbon fractionsboiling primarily in the gasoline or naphtha range, in a multiple,adiabatic, fixed bed catalyst conversion system employing a supportedplatinum group metal catalyst and a catalyst containing a platinum groupmetal and rhenium on a solid, acidic oxide support to improve the octanenumber of the feed.

In view of the endothermic nature of catalytic reforming reactions thereis usually employed a series of adiabatic catalyst bed reactors in suchoperations. One method comprises preheating the naphtha charge to thedesired inlet temperature, passing it to the first reactor, reheatingetfiuent from the first reactor and passing it into the second reactor,and so on through the remaining reheaters and reactors of the series.The inlet temperatures of each of the reactors can be the same ordifferent and they generally fall in the range of about 850 to 960 F. ormore. The temperature drop exhibited in each of the series of catalystbeds progressively decreases in the direction of hydrocarbon flow. Forinstance, the temperature drop in the first reactor of a three reactorseries usually ranges between about 50 to 150 F., whereas thetemperature drop in the terminal reactor is generally about 25 F.maximum and the last reactor may be exothermic, especially in highoctane-higher pressure operations, e.g., 325 p.s.ig. or more. In thesesystems it is generally considered that in one or more of the reactorsin the first part of the series, the predominant reaction isdehydrogenation, for instance, the conversion of naphthenes toaromatics, while in one or more of the reactors in the terminal part ofthe series a principal reaction is paratfin dehydrocyclization. It isalso known that catalysts composed essentially of small amounts of aplatinum group metal or both a platinum group metal and rhenium on asolid support can be used in such reforming operations.

The present invention is based on the use in at least one of thenaphthene dehydrogenation zones of such reforming systems, of a catalysthaving a platinum group nited States Patent metal as the essentialpromoting metal component, for instance, the catalyst contains nosignificant amount of a rhenium component; and the use in at least oneof the parafiin dehydrocyclization zones of a catalyst having both aplatinum group metal and rhenium as the essential catalytic metals on aporous, solid, acidic oxide support. When the reforming operation isconducted with these distinct catalysts in the separately designatedreaction zones a highly advantageous yield-octane number-catalyst agingrelationship is obtained.

In the operation of this invention the naphthene dehydrogenationreaction zones serve to convert naphthenes to aromatics and perhapsthere is also a minor amount of parafiin dehydrogenation. In a givensystem there may be only a single naphthene dehydrogenation reactionzone, although the series of catalyst beds may have 2, 3 or even 5 ofsuch zones with the number generally rising as the naphthene content ofthe feed increases. In the present invention the naphthenedehydrogenation zones employing the designated catalyst have an inlettemperature of at least about 820 F. and it is preferred that thesezones include the first catalyst bed of the series. Often the inlettemperatures for the naphthene dehydrogenation zones will be in therange of about 820 to 920 F., preferably about 840 to 890 F. In somereforming operations there may be provided a preliminary reactor inwhich the inlet temperature is less than about 820 F. In some reformingoperations there may be provided a preliminary reactor in which theinlet temperature is less than about 820 F. and naphthenedehydrogenation may take place in this reaction zone. In such systemsthe heating of the feedstock is usually accomplished by indirectexchange, for instance, with the reformate product, rather than by afired preheater. In the method of the present invention the designatedplatinum group metal catalyst having essentially no rhenium component isemployed in at least the initial, and preferably all, of the naphthenedehydrogenation zones which have inlet temperatures of at least about820 F. The catalyst in any lower temperature reactor may be selectedaccording to the desires of the operator, but again is preferably aplatinum group metal catalyst having essentially no rhenium component.By the term inlet temperatures, reference is made herein to thetemperature of the initial portion of the catalyst in the reactor inquestion.

In a preferred operation the inlet temperatures to the naphthenedehydrogenation zones are in the range of about 820 to 920 F. for atleast about of the total reforming processing cycle. When there are atleast two such naphthene dehydrogenation reactors, it is preferred thatthe inlet temperature of the first such reactor not exceed about 900 F.while the inlet temperature of the subsequent naphthene dehydrogenationreactors be maintained below about 920 F., for at least about 80% of thetotal reforming processing time. As a further refinement in this type ofoperation the total temperature drop in each of the naphthenedehydrogenation zones does not vary more than about 30 F. over at leastabout 80% of the total reforming cycle. The reforming system can be runso that there is about 75 to Weight percent conversion of naphthenes inthe overall dehydrogenation zones which provides an effluent from thelast-in-series naphthene dehydrogenation reaction zone having less thanabout 10 weight percent naphthenes. Frequently, the conditions in thenaphthene dehydrogenation zones include pressures of about 50 to 500p.s.i.g., preferably about to 350 p.s.i.g., and weight hourly spacevelocities for the overall dehydrogenation reaction zones of about 0.5to 4 WHSV.

As previously stated, at least one of the reactors in the terminal partof the series, and preferably at least the last, employs a catalystcontaining a platinum group metal and rhenium supported on a solid,acidic oxide base and a principal reaction effected is thedehydrocyclization of paraflins to aromatics. There may be more than oneof such parafiin dehydrocyclization reactors and each of such reactorshas an inlet temperature of about 900 to 1000 F., preferably about 900to 960 F. It is further preferred that such inlet temperatures be atleast about 20 F. greater than the inlet temperatures of any of thenaphthene dehydrogenation reactors for at least about 50% of the totalreforming process time. Frequently, the inlet temperature of the lastreactor of the series exceeds about 930 F. for at least about 25% of thetotal reforming cycle. The catalyst volume in the total naphthenedehydrogenation zones compared to that in all of the paraffindehydrocyclization zones is often about 1:20 to 3:1; preferably thisratio is at least about 1:5 when the naphthene content of the gasolineor naphtha feed is at least about 30 volume percent, and the ratio isless than about 1:4 when the paraffin content of the gasoline or naphthafeed is at least about '60 volume percent. Although the system of thisinvention is described with reference to naphthene dehydrogenationreaction zones and paraffin dehydroeyclization reaction zones there maybe other catalytic reactors in the series which are operated underconditions such that they do not fall in the category of the describednaphthene dehydrogenation or paraffin dehydrocyclization zones.

During the reforming operation hydrogen is supplied to the reactionzones by recycle .of hydrogen-containing gases separated from theproduct effiuent. The recycle gas in the method of this invention isrelatively high in hydrogen content and low in methane and thereforeaffords a higher hydrogen partial pressure for a given total pressureand aids in extending catalyst life. Also the purer hydrogen can be morereadily used in other processes.

Although the total hydrogen gas stream which is recycled can be passedto the initial naphthene dehydrogenation reaction zone and thencethroughout the series of reactors, it is preferred that only a portionof the recycle gases go to the naphthene dehydrogenation reaction zoneswith there being separate hydrogen gas introduction to the paraffindehydrocyclization zones. Thus, there may be provided about 0.5 to 8moles, preferably about 1 to 4 moles, of hydrogen recycle gas per moleof hydrocarbon boiling in the gasoline range introduced into the initialnaphthene dehydrogenation zone which has an inlet temperature of about820 F., while there is passed directly to the parafiindehydrocyclization Zones sufficient hydrogen gas to give in such zonesabout 7 to 30 moles, preferably about 8 to 15 moles, of total hydrogenrecycle gas per mole of hydrocarbon boiling in the gasoline rangepassing into the paraffin dehydrocyclization zones. Of course the totalgas to such dehydrocyclization zones includes that coming from thepreceding reactors as well as any portion of the recycle gas passeddirectly into the paraffin dehydrocyclization zones. The latteraddition, that is, the direct charging of recycle gas to the parafiindehydrocyclization zones, is usually at least about 3 moles of hydrogenrecycle gas, preferably at least about moles, per mole of hydrocarbonboiling in the gasoline range passing into such zones.

As noted before one of the catalysts employed in the process of thepresent invention contains as promoters small, catalytically-effectiveamounts of each of a platinum group metal and rhenium supported by aporous, acidic oxide base. The base has a minimum activity of at leastabout 20 D+L as measured by the method of Birkhimer et al., A BenchScale Test Method for Evaluating Cracking Catalysts, Proceedings of theAmerican Petroleum Institute Division of Refining, vol. 27 (III), page90 (1947), preferably such activity is about 20 to 40 D+L. Surface areasof such bases are usually a minimum of about 100 square meters per gram,preferably about 150 to 500 square meters per gram. The supportconstitutes the major portion of the catalyst and can be of a variety ofmaterials among which are combinations of silica, alumina, zirconia,titania, magnesia, boria and aluminosilicates, especially thosecrystalline aluminosilicates having relatively uniform pores havingopenings whose diameters are about 6 to 15 angstrom units, preferably inthe approximate 10 to 14 angstrom unit range. These supports, which arepreferably metal oxides, can also contain halogens, especially fluorineand chlorine.

Particularly useful acidic materials are silica-alumina, includingaluminosilicates, especially crystalline aluminosilicates. Thecrystalline aluminosilicates usually have silica to-alumina mole ratiosof at least about 2:1, for instance about 2 to 12:1, preferably about 4to 6:1. The crystalline aluminosilicates are usually available or madein sodium form and this component can be reduced, for instance to lessthan about 4, or even less than about 1, weight percent, through ionexchange with hydrogen ions, hydrogen-precursors such as ammonium ions,or polyvalent metals. Suitable metals include the rare earths such ascerium, and their mixtures. Mixtures of alumina and amorphoussilica-alumina cracking catalysts, especially those having a majorproportion of silica, e.g. about 60 to Weight percent silica and about10 to 40 weight percent alumina, are suitable bases, Advantageous basesare composed of mixtures of alumina and crystalline aluminosilicates,for instance containing about 0.1 to 25 weight percent of thecrystalline aluminosilicate, preferably about 1 or even 5 to about 15weight percent. The alumina in such mixed bases can be of the typedescribed below with respect to the non-rhenium catalysts.

The platinum group metal and rhenium in the foregoing described acidicsupport catalysts are each often about 0.05 to 3 Weight percent of thecatalyst, preferably about 0.1 or 0.3 to 1 weight percent. Platinum isthe most preferred metal in such catalysts but other platinum groupmetals such as palladium and rhodium can be used. When the catalyst isin a virgin state the promoting metals are preferably for the most partundetectable by X-ray diffraction analysis, which indicates that if themetals are present in the catalyst as elemental metals or alloys theircrystallite sizes are less than about 50 A. A common method of providingthe platinum group metal in the catalyst is by contact of one or more ofthe porous, acidic oxide components of the support, either in hydrous orin calcined form, with an aqueous solution of a chlorine-containingcompound, for instance chloroplatinic acid. In this manner chlorine isincorporated in the catalyst for instance, in amounts of about 0.2 to 2weight percent, preferably about 0.3 to 1 weight percent. Such amountsof chlorine component can also be provided in the catalyst from a sourceother than the compound supplying the platinum group metal. Rhenium canalso be added to one or more of the porous, acidic oxide components ofthe hydrous or calcined support through contact with an aqueous solutionof a rhenium compound, e.g. perrhenic acid or ammonium perrhenate. Ifpromoting metal is added to a calcined support, the resulting compositeis generally recalcined.

The catalyst which can be used in one or more to all of the naphthenedehydrogenation zones and even in all but one paraffindehydrocyclization reactor of the system, contains a platinum groupmetal on a support and has an essential absence of rhenium, e.g. lessthan about 0.05 Weight percent, or less than about 0.01 weight percentor even no detectable amount, of rhenium. As stated above, it ispreferred that these catalysts, which can be designated as non-rheniumcatalysts, be used in all reaction zones except the paraffindehydrocyclization zones. These non-rhenium catalysts can be made byvarious procedures, for instance those described above with respect tothe platinum group metal and rhenium containing catalysts, but, ofcourse, without any significant addition of rhenium and with the use ofrelatively non-acidic, porous solid supports. The platinum group metalis present in a small, promoting amount sufficient to provide reformingor de hydrogenation activity to the catalyst. Such amounts often includeat least about 0.05 weight percent of the platinum group metal,preferably about 0.1 to 0.3 to 1%, based on the total weight of thecatalyst. The major portion and essential balance of the catalyst is therelatively non-acidic support, although the catalyst can contain minoramounts of other promoters as long as the rhenium content is notsignificant and the activity of the catalyst is not increased toogreatly. Thus the support or base, with any other non-platinum groupmetal component has an activity below about 15 D-l-L as measured by theabove-mentioned test, preferably below about D+L. Alumina is thepreferred non-acidic support, and platinum is the preferred platinumgroup metal in the catalyst.

The alumina support in the non-rhenium catalyst of the present inventionoften has a surface area of at least about 150 square meters per gramand is preferably composed to a major extent of gamma-family aluminamodifications derived by the activation or calcination of aluminatrihydrates. These gamma-family or activated alumina modificationsinclude among others, gamma and eta aluminas. US. Pat. No. 2,838,444discloses this type of alumina support having surface areas in the rangeof about 350 to 550 square meters per gram, while in US. Pat. No.2,838,445 there is described catalyst supports made from predominantlytrihydrate alumina precursors, the supports having surface areas in therange of about 150 to 350 square meters per gram. These supports aresuitable for use in the present invention, especially the higher areasupports of Pat. 2,838,444 which supports during use may have theirsurface areas reduced to about 150 to 250 square meters per gram. Asstated, the preferred alumina precursors predominate in trihydrate whichmay contain one or more of the bayerite, gibbsite or nordstrandite(previously called randomite) forms, and prefera'bly a major amount ofthe trihydrate is composed of bayerite or nordstrandite which whencalcined can form eta alumina. It is also advantageous that the hydrousalumina precursor contain about 65 to 95% of the trihydrate with theessential balance being composed of one or both of the aluminamonohydrate, boehmite, or amorphous hydrous alumina. Preferred supportshave pore volumes of at least about 0.1 cc./gm.; preferably at leastabout 0.15 cc./gm., in pores greater than about 100 A. radius. It isalso preferred that the supports have at least about 0.05 cc./gm. inpores greater than about 300 A. or even greater than about 600 A.radius. These determinations are by the method described by Barrett,Joyner and Halenda, JACS, 73, p. 373 (1951).

Calcination of both types of catalysts used in this invention can beconveniently conducted at temperatures of the order of about 700 to 1200F. or more, for instance in an oxygen-containing gas, and this operationcan be controlled to give a final catalyst of desired surface area. Atan appropriate stage in the manufacture of the catalysts, the particlescan be formed into macrosize as distinguished from finely divided orfluidized catalyst types. The macrosize particles frequently havediameters in the range of about to inch, preferably about to 4 inch, andif not spherical, the particles usually have lengths of about ,4 to 1inch or more, preferably about A; to /2 inch.

The process of this invention involves hydrocarbon reforming conductedat elevated temperatures up to about 1000 F., and under a reducingatmosphere provided by the presence of a molecular hydrogen-containinggas. The feedstocks include gasoline boiling range hydrocarbons whetherthey boil over a broad or narrow temperature range. In such operations anaphtha, a fraction thereof or other similar boiling range hydrocarbonswhose aliphatic and cycloaliphatic constituents are for the most partsaturated and which may contain some aromatics, is converted to aproduct having greater aromaticity and higher octane rating. Relativelypure aromatics can be separated from the products. The feeds employed inthe process of the invention include naphthas composed of at least about15% up to about 7 0% (by volume) naphthenes and at least about 25%parafiins and generally have clear or unleaded research octane ratings(RON) in the range of about 30 to 60. Advantageously, the totalhydrocarbon feed and recycle gas passing to a given reactor of theinvention contain less than about 10 p.p.m. (by weight) sulfur, andpreferably less than about 5 p.p.m. combined nitrogen. Superioroperations can be provided where the hydrocarbon feed and recycle gasstream to all reactors have less than about 5 p.p.m. sulfur, and lessthan about 2 p.p.m. combined nitrogen. These impurity levels are basedon the weight of total process materials passing to a given reactor.

Although the reforming system is advantageously operated with less thanabout 30 p.p.m. by volume Water, preferably less than about 10 p.p.m.,based on the hydrogen gas present, it can be enhanced by the provisionof controlled amounts of water in the paraffin dehydrocyclizationreactors, especially where the catalyst base contains amorphoussilicaalumina as the essential cracking component. Thus we can supplyabout 15 to 250 p.p.m. H O, preferably about 20 to 100 p.p.m., to thedehydrocyclization zones to improve the operation While drying therecycle gas to less than about 10 p.p.m. H O for passage to thenaphthene dehydrogenation zones.

The reforming systems are usually conducted at processing conditionswhich include reactor inlet temperatures of about 825 to 975 or 1000 F.,and total pressures of about 50 to 600 p.s.i.g., preferably about 100 to350 p.s.i.g. During the operation hydrogen-containing eflluent gas isrecycled to the reaction system, the latter having a series of adiabaticfixed bed catalyst reactors each being preceded by a feed heater. Therecycle gas ratio is usually such to provide about 3 to 30 moles ofhydrogen gas per mole of gasoline boiling range hydrocarbon feedstock.Also the hydrocarbon charge is often passed to the reactor system at arate such that the overall space velocity .is about 0.5 to 15 WHSV(weight of hydrocarbon per weight of catalyst per hour) preferably about1 to 6 WHSV. The severity of the reaction conditions are such that thenormally liquid reformate or product from the terminal reactor has a RONof at least about or even at least about or 100. Carbonaceous depositsaccumulate on the catalysts of this invention as reforming proceeds andas a result the catalysts lose activity which can be counteracted byincreasing the reaction temperature. Eventually however, when thereactor inlet temperatures reach a desired maximum, for instance in therange of about 950 to 1000 F., especially at about 970 F. and above, itbecomes inadvisable to increase the temperature further, otherwise undueaging of the catalysts may result. The catalysts can then be regeneratedby carbon burn-off which improves the catalytic characteristicssufficiently for the catalysts to be reused on an economic basis.

At the beginning of regeneration the carbon content of the catalysts isgenerally above about 0.5 Weight percent, often greater than about 10weight percent. During regeneration of the catalysts by burning, thecarbon level is often reduced to below about 0.5 weight percent,preferably below about 0.2 weight percent. This burning is conductedthrough contact of the catalysts with an oxygen-containing gas andgenerally the amount of oxygen is controlled to maintain the temperatureof the catalysts from about 700 to about 900 or 1000 -F., preferably inthe temperature range of about 700 to 850 F. The pressure maintainedduring burning is preferably elevated, for instance, is about 50 to 5 00p.s.i.g. The controlled burning is usually initiated with an inert gas,e.g. nitrogen, carbon dioxide or their mixtures, containing a smallamount of oxygen, for instance, up to about 1 mole percent andpreferably With an oxygen partial pressure of at least about 0.2p.s.i.g. When the bulk of the carbon has been removed from the catalystsby a gas containing the relatively low concentration of oxygen, theamount of oxygen can be increased somewhat to insure that sufiicientcarbon has been removed from the catalysts without exceeding the desiredtemperature. This type of treatment is exemplified by one or moreburns-through of the catalyst bed at about 800 F. to 850 F., and about100 to 500 p.s.i.g., with a gas containing above about 0.5 to about 3 orsomewhat greater mole percent oxygen. Other suitable carbon-burningprocedures can be employed as long as the temperatures are controlledand the carbon level of the catalysts is adequately lowered. Duringcarbon burnoff and subsequent treatments of the catalysts with anoxygen-containing or other gas at elevated temperatures, the gas shouldbe dry enough to avoid undue additional sintering of the catalysts andloss of surface area. Such loss generally increases as temperature,water content of the gas or treating time is raised.

Especially where the crystallite size of the promoting metals on thecatalysts is to be reduced, the catalysts can, after carbon burn-off, becontacted with an oxygencontaining gas at a temperature of about 800 to1000 F., preferably about 850 to 950 F., and if desired, an elevatedpressure such as about 100 to 500 p.s.i.g. This treatment has sometimesbeen referred to in the art as an air soak and the oxygen content of thegas is usually greater than that present in the gas employed for carbonburn-off. Thus, the oxygen content of the gaseous stream employed forair soaking is often at least about 5 mole percent with there havingbeen found no particular reason for increasing the gas content aboveabout 20 mole percent. The air soaking period is generally at leastabout one hour and is usually continued for several hours, for instance,in the range of about 5 to 24 hours. Regeneration and air soakingprocedures suitable for the catalysts of the present invention aredisclosed in US. Pat. No. 2,922,756, herein incorporated by reference.

The virgin catalysts of this invention or used catalysts, of such types,say after regeneration with or without reactivation, can be reduced bycontact with a gaseous stream which contains molecular hydrogen. Thetreatment is at an elevated temperature, for instance, about 600 to 1000F., preferably about 750 to 950 F. Elevated pressures are preferablyused in the reduction and can be, for example, about 20 to 600 p.s.i.g.,preferably about 50 to 350 p.s.i.g. Apparently, the reduction convertsthe catalytic promoting metals to their elemental state, but if avaporous sulfiding agent be present some or all of the promoting metalsmay be sulfided. By using an essentially dry, hydrocarbon-free gasduring the reduction, hydrocracking is avoided as are its attendantdisadvantages of, for instance, excessive catalyst temperature rises andthe formation of catalyst poisons or deleterious agents such as carbonmonoxide which can cause undesirable crystallite growth of the catalyticpromoting metals. Also, carbon monoxide, for example, can interact withthe catalytic promoting metals causing deactivation. The gas streamemployed during reduction is often composed of about 70 to 100 volumepercent hydrogen, preferably about 95 or 99 to 100 volume percent, withany remaining components being up to about 30 volume percent of inertgas such as nitrogen. The gas advantageously is dry and contains lessthan about 1 volume percent hydrocarbons, preferably less than about0.1%.

To avoid undue hydrocracking of the hydrocarbon feedstock during theinitial period of hydrocarbon processing after the catalysts of thepresent invention, especially the platinum group metal-rhenium catalyst,are placed on-stream, the catalysts can be contacted with a gascontaining sulfur-providing component in vaporous form. This sulfidingtreatment can be conducted simultaneously with or subsequent to thereduction. If sulfiding is conducted simultaneously with the reduction anoncarbonaceous sulfur compound should be used due to the presence ofoxygen in the system and to avoid any localized overheating of thecatalyst. Suitable sulfurproviding materials or sulfiding agents includeS0 and H 3, preferably the latter. The amount of sulfiding agentemployed is at least about 25% or even at least about of thestoichiometric amount needed to give one atomic weight of sulfur foreach atomic weight of total platinum group metal and rhenium in thecatalyst, preferably the amount is at least about 50% to say up to about500% or more. The sulfiding operation can be done at an elevatedtemperature, e.g. about 650 to 950 F., and at any suitable pressure,preferably an elevated pressure such as about 100 to 500 p.s.i.g. Thesulfiding gas is reductive and usually contains a minor amount of thesulfur-bearing component, e.g. about 01-10 volume per cent, preferablyabout 0.2 to 3%, with the major component being hydrogen or an inert gassuch as nitrogen. Also the sulfiding agent can be added to the inlet ofeach reactor of the hydrocarbon processing system to m1n1- mize contactwith other equipment surfaces where corrosion might occur. When thesulfiding is conducted simultaneously with or subsequent to reducing thecatalysts with hydrogen, the catalysts are in sulfided form when theyfirst contact the hydrocarbon being processed which avoids excessivehydrocracking with its attendant yield and selectivity losses.

It can be further advantageous in minimizing hydrocracking caused by thereduced catalysts whether presulfided or not, to supply vaporoussulfiding agent to the conversion system when charging of thehydrocarbon feedstock is begun. Thus, a small amount of the sulfidingagent, sutficient to significantly reduce hydrocracking during theinitial portion of the processing cycle, can be added to the system. Thesulfiding agent can conveniently be charged with the recycle gas or withthe hydrocarbon stream. The amounts of sulfiding agent employed includeabout 1 to 500 p.p.m. by volume based on the hydrogen passing to thereaction system, preferably about 5 to 200 p.p.m. This sulfiding-agentaddition can be continued as long as the operator desires but often theaddition will approximate the time period in which, in the absence ofthe sulfiding-agent addition, the catalysts would cause significantlyexcessive hydrocracking. Hydrocracking can be detected in the processingsystem by any desirable means such as a drop in the hydrogen content ofthe offgases, a change in the ratio of methane to propane in the gases,or the temperature rise in the catalyst beds. The period ofsulfiding-agent addition upon placing the reduced catalysts back onprocessing can include, for instance, about 1 to or more days and isoften about 3 to 10 days.

The present invention will be further illustrated by the followingexample.

A 20,000 BPSD naphtha reforming operation is carried out in a fourreactor system employing a straight run naphtha feed containingapproximately 42% naphthenes, approximately 43% parafiins andapproximately 15% aromatics. The naphtha feed has a RON of 39, a boilingrange of 145 F. to 385 F., less than about 5 p.p.m. H O, less than about4 p.p.m. S, less than about 2 p.p.m. N, and less than about 1 p.p.m. C].The naphtha feed in admixture with 3 mols of hydrogen-containing recyclegas is heated to temperatures of about 890 F., and passed to the initialreactor. The reactor system is at a pressure of about 250 p.s.i.g. Thefirst reactor, as well as each of the next two reactors, contains afixed bed of catalyst having about 0.6% platinum on alumina (5 D+Lactivity). The fixed bed catalyst in the last reactor contains about0.6% Pt and 0.6% Re on a base (30 D+L activity) composed of 90% aluminaand 10% of a hydrogen-exchanged, faujasite-type crystallinealuminosilicate having a silica-to-alumina mole ratio of about 4.5 :1and pore openings of about 13 A. size.

Both catalysts are extrudates, and the catalyst in the first threereactors has an alumina base of the type disclosed and claimed in US.Pat. 2,838,444. The alumina is derived from a hydrous alumina mixturehaving about trihydrate which is predominantly bayerite andnordstrandite, the essential balance of the hydrous alumina beingboehmite and amorphous hydrous alumina. Both catalysts in the calcined,virgin state have surface areas above about 400 square meters per gram,and chloride contents of about 0.7 weight percent derived from the useof chloroplatinic acid as the source of platinum.

The catalyst-reactor distribution is 1:l:1:3 by volume and the overallweight hourly space velocity is 2.0. The effluent from the first reactorwhich is at a temperature about 125 F. lower than the inlet temperatureis heated to 900 F. for charging to the second reactor. The temperatureof the eflluent from the second reactor is about 100 F. lower than theinlet temperature, and the eflluent is heated to 920 F., prior to beingintroduced into the third reactor. Eflluent from the third reactor is ata temperature about 50 F. lower than the inlet temperature of thereactor. This elfiuent containing about 5% by weight naphthenes, iscombined with the remainder of the hydrogen-containing gas. The combinedfeed and recycle gas are then heated to a temperature of 955 F., priorto introduction into the fourth reactor where the reforming iscompleted. The naphthene content of the third reactor effluent of about5% represents a conversion of naphthenes to aromatics in the first threereactors of about 80%. Hydrogen and light gases, including hydrocarbongases such as methane, ethane and propane are separated from the C+liquid reformate of 99 Research Octane Number (clear). A portion ofhydrogen and light hydrogen and light hydrocarbon gases is recycled andcan be treated for removal of sulfur, nitrogen and water. The gas isthen pressured to about 300 p.s.i.g. The repressured recycle gas issplit into two streams on the basis of about 3 moles of gas per mole offresh naphtha for mixing with the incoming fresh naphtha. The secondstream of recycle gas is combined with the third reactor efiluent at arecycle rate of about 8 moles of gas per mole of naphtha feed (totalrecycle: 11 moles of gas per mole of feed).

During the processing cycle the inlet temperature of the last reactor isperiodically raised to maintain a yield of 99 RON reformate. The inlettemperatures of the first three reactors are periodically raised by F.,10 F. and 5 F., respectively, to prevent the change in the temperaturedrop in each of these reactors from varying more than approximately 7 F.in each reactor. Typical inlet temperatures of the reactors. at the endof the processing cycle of the respective reactors are 910 F., 920 F.,935 F., 1000 F., for the first, second, third and fourth reactors,respectively.

It is claimed:

1. In a method of reforming naphthene and parafiincontaining petroleumhydrocarbons of gasoline or naphtha boiling range in the presence ofmolecular hydrogen wherein is employed in series a plurality ofcatalytic reaction zones to provide reformates of at least about 90 RONand hydrogen-containing recycle gas, the improvement which comprisesproviding at least one naphthene dehydrogenation zone in the earlyportion of said plurality of reaction zones, at least the initial saidnaphthene dehydrogenation zone having a catalyst consisting essentiallyof a platinum group metal on a relatively non-acidic support and havingan essential absence of rhenium, the support of said catalyst having aD+L activity below about 15, and at least one parafiindehydrocyclization zone in a latter portion of said plurality ofreaction zones, paraffin dehydrocyclization zone having a catalystconsisting essentially of a platinum group metal and rhenium on a solid,porous, acidic oxide support having a D+L activity of at least about 20,introducing hydrocarbon of gasoline or naphtha boiling range into saidnaphthene dehydrogenation zone at an inlet temperature of at least about820 F., passing efiluent from said naphthene dehydrogenation zone tosaid paraffin dehydrocyclization zone,

the inlet temperature of such parafiin dehydrocyclization being about900 to 1000 F. and run to give a reformate of at least about 90 RON.

2. The method of claim 1 wherein the support of the catalyst in the saidinitial naphthene dehydrogenation zone is alumina.

3. The method of claim 2 wherein the support of the catalyst in theparaffin dehydrocyclization zone is comprised of silica and alumina.

4. The method of claim 3 wherein the platinum group metal of thecatalysts is platinum.

5. The method of claim 3 wherein the alumina of the catalyst in saidinitial naphthene dehydrogenation zone is derived by calcination ofhydrous alumina predominating in trihydrate.

6. The method of claim 5 wherein the platinum group metal of thecatalysts is platinum.

7. In a method of reforming naphthene and parafiincontaining petroleumhydrocarbons of gasoline or naphtha boiling range in the presence ofmolecular hydrogen and supported platinum group metal reformingcatalysts wherein is employed in series a plurality of catalyticreaction zones, each of said plurality of zones being preceded byheating means for the hydrocarbon processed and molecular hydrogen, toprovide reformates of at least about 90 RON and hydrogen-containingrecycle gas, the improvement which comprises providing at least onenaphthene dehydrogenation zone in the early portion of said plurality ofreaction zones, at least the initial said naphthene dehydrogenation zonehaving a catalyst consisting essentially of a platinum group metal on asupport and having an essential absence of rhenium, the support of saidcatalyst having a D+L activity below about 10, and at least one paraflindehydrocyclization zone in a latter portion of said plurality ofreaction zones, said paraffin dehydrocyclization zone having a catalystconsisting essentially of a platinum group metal and rhenium on a solid,porous, acidic, metal oxide support, intro ducing said hydrocarbon ofgasoline or naphtha boiling range containing at least about 15% byvolume of naphthenes and at least about 25% by volume of paraflins intosuch naphthene dehydrogenation zone at an inlet temperature of about 820to 920 F. for at least about of the total reforming process time whilepassing a portion of said recycle gas to such naphthene dehydrogenationzone at a rate of about 0.5 to 8 moles of recycle gas per mole ofgasoline boiling range hydrocarbon feed and for a reaction timesufficient to provide an efiluent from the last such naphthenedehydrogenation zone having less than about 10% by weight naphthenes,passing efiiuent from said naphthene dehydrogenation zone to saidparaflin dehydrocyclization zone, the inlet temperature of such parafiindehydrocyclization zone being about 900 to 1000 F. and run to give areformate of at least about RON, said dehydrocyclization zone inlettemperature being at least 20 F. greater than the inlet temperature ofthe first naphthene dehydrogenation zone for at least about 50% of thetotal reforming process time, while passing a portion of thehydrogen-containing recycle gas to such parafiin dehydrocyclization zoneat a rate such that the total gas recycle to the paraffindehydrocyclization zone is about 7 to 30 moles of said recycle gas permole of gasoline boiling range feed, said portion of hydrogen-containingrecycle gas to said paraffin dehydrocyclization zone being at least athird of the total hydrogen-containing recycle gas recycled, thecatalyst volume distribution of the naphthene dehydrogenation zones tothe paraflin dehydrocyclization zones being between about 1:20 to 3:1,and maintaining such zones under endothermic conditions.

8. The method of claim 7 wherein the platinum group metal of thecatalysts is platinum.

9. The method of claim 8 wherein the support of the catalyst in thenaphthene dehydrogenation zone is alumina.

11 10. The method of claim 9 wherein the support of the catalyst in theparaffin dehydrocyclization zone is comprised of silica and alumina.

11. The method of claim 10 wherein the inlet temperature of thenaphthene dehydrogenation zone is about 840 to 890 F.

12. The method of claim 11 wherein the mole ratio of hydrogen-containingrecycle gas to gasoline boiling range hydrocarbon in the naphthenedehydrogenation zone is about 1 to 4:1 and the mole ratio ofhydrogen-containing recycle gas to gasoline boiling range hydrocarbon inthe paraffin dehydrocyclization zone is about 8 to 15:1.

13. The method of claim 1 wherein the sulfur content of the totalhydrocarbon feed and recycle gas passing to said naphthenedehydrogenation zone is less than about 10 p.p.m.

14. The method of claim 13 wherein the platinum group metal of thecatalysts is platinum.

15. The method of claim 14 wherein the impurity levels of both thehydrocarbon feed and recycle gas are below about 10 p.p.m. water, 5p.p.m. sulfur, and 2 p.p.m. combined nitrogen.

16. The method of claim 15 wherein the support of the naphthenedehydrogenation zone catalyst is alumina derived by calcination ofhydrous alumina predominating in trihydrate.

17. The method of claim 7 wherein the sulfur content of the totalhydrocarbon feed and recycle gas passing to said naphthenedehydrogenation zone is less than about 10 p.p.m.

18. The method of claim 17 wherein the platinum group metal of thecatalysts is platinum and the support of the catalyst in the naphthenedehydrogenation zone is alumina.

'19. The method of claim 1 wherein the platinum group metal of thecatalysts is platinum and the support of the catalyst in the naphthenedehydrogenation zone is alumina.

20. The method of claim 2 wherein the platinum group metal of thecatalysts is platinum and the support of the catalyst in the naphthenedehydrogenation zone is alumina.

21. The method of claim 1 wherein about 15 to 250 p.p.m. of H 0 aresupplied to said paraffin dehydrocyclization zone.

22. The method of claim 3 wherein about 15 to 250 p.p.m. of H 0 aresupplied to said paraffin dehydrocyclization zone.

23. The method of claim 4 wherein about 15 to 250 p.p.m. of H 0 aresupplied to said paraffin dehydrocyclization zone.

24. The method of claim 23 wherein the sulfur content of the totalhydrocarbon feed and recycle gas passing to said naphthenedehydrogenation zone is less than about 10 p.p.m.

References Cited UNITED STATES PATENTS 3,287,253 11/1966 McHenry et al.208-65 3,415,737 12/1968 Kluksdahl 2O8139 3,436,335 4/1969 Maziuk 208-652,838,444 6/1958 Teter et al 2'08--138 2,838,445 6/1958 Teter et a1.208l38 HERBERT LEVINE, Primary Examiner US. Cl. X.R. 208138, 139

